Hydrocarbon conversion process includ-
ing a hydrocracking stage, two stages
of catalytic cracking, and a reform-
ing stage



KING

March 9, 1965 N. J. PATERSON HYDROCARBON CONVERSION PROCESS INCLUDING A HYDROCRAC STAGE, Two STAGES oF CATALYTTC CRACKTNG, AND

A REFORMING STAGE Filed March 30, 1962 United States Patent 3,112,842 HYDROCARBON CNVERSEON PRUCESS MCLUD- ING A HYDROCRACKMG STAGE, TWO STAGES F CATALYTEC CRACKHQG, A REFORM- ING STAGE Norman Il. Paterson, San Rafael, Calif., assignor to California Research Corporation, San Francisco, Calif., a corporation of Delaware Filed Mar. 30, 1962 Ser. No. 183,972 4 Claims. (Cl. 208-80) INTRODUCTION This invention relates to a process for the catalytic conversion of petroleum hydrocarbons, including gas oil and residual portions thereof, to gasoline and middle distillate fractions. More particularly, the invention relates to an integrated refinery process wherein nonresidual and residual hydrocarbon stocks boiling essentially above the gasoline range are converted to fuel values, including high octane gasoline and middle distillate fractions, in an exceptionally high liquid yield.

DEFINITIONS Because throughout this specification numerous terms will be used to characterize various hydrocarbon charge stocks and fractions thereof, various conversion products, and various characteristics of the aforesaid stocks, fractions and products, such terms will rst be defined in order to facilitate understanding of the subsequent description.

In order to avoid thermal cracking, the distillation of petroleum fractions boiling above about 750 F. conventionally is carried out under vacuum. However, throughout the following description the boiling ternperatures given refer to the boiling point at atmospheric pressure, i.e., for uniformity with the boiling points referring to atmospheric pressure distillations, the boiling points referring to vacuum distillations have been corrected to the corresponding boiling points at atmospheric pressure.

The process of the present invention is capable of converting an entire petroleum crude hydrocarbon feed to fuel values in an exceptionally high liquid yield, and is particularly effective in so converting a paraffinic crude, for example a Mid-Continent, Middle East, Canadian, West Texas or East Texas crude, although it also is applicable to oils recovered from shale, gilsonite, tar sands and the like. The term crude refers to the crude petroleum as recovered from an oil well, after separation therefrom of constituents that are gaseous under recovery conditions at the Well, for example constituents such as components which are gaseous at recovery conditions but which are condensed to liquid products.

Upon distillation, various fractions may be separated from the crude. A fraction which may boil from about C5 to about 180 F. is termed a light gasoline (also known as light naphtha). A fraction which boils from about 180 to about 400 F. is termed a heavy gasoline (also known as heavy naphtha), and is conventionally used as a reformer charge stock.

After the light gasoline and heavy gasoline have been distilled from the crude, the entire remaining portion of the crude is called a whole topped crude.

The fractions distilling off after the gasolines are called distillates or gas oils. It is well known that the term gas oil is a broad, general term that covers a variety' of stocks. When used in the following description, the term, unless further modified, includes any fraction distilled from petroleum which has an initial boiling point of at least about 350 F., a 50% point of at least about 475 F. and an end point of at least about 600 F., and boiling substantially continuously between "C ld ce Patented Mar. 9, 1955 the initial boiling point and the end point. The portion of the crude oil which is not distilled is considered to be a residual stock or residuum The exact boiling range of a gas oil therefore will be determined by the initial distillation temperature (initial boiling point), the 50% point, and by the temperature at which distillation is cut off (end point).

ln practice, petroleum distillations have been made under vacuum up to temperatures as high as ll00 to l200 F. (corrected to atmospheric pressure). Accordingly, in the broad sense a gas oil is a petroleum fraction which boils substantially continuously between two temperatures that establish a range falling Within from about 350 to about 1l00 to l200 F., the 50% point being at least about 475 F. Thus, a gas oil could boil over the entire range of about 375 to 1200 F., or it could boil over a narrower range, for example 500 to 900 F.

The gas oils can be further roughly subdivided by overlapping boiling ranges. Thus, a light gas oil boils between about 375 and 650 F. A medium gas oil boils between about 600 and 750 F. A heavy gas oil boils between about 600 and 900 F. A gas oil boiling between about 800 and 1200 F. is sometimes designated as a vacuum gas oil. It must be understood, however, that a gas oil can overlap the foregoing ranges. It might even span several ranges; for example, it may include both light and medium gas oils.

As heretofore mentioned, a residual stock or residum is any portion which is not distilled. Therefore, any portion, regardless of its initial boiling point, which includes all the heavy bottoms such as tars, asphalte, etc., is a residual stock. For example, a residual stock can be the portion of the crude remaining undistilled at l200 F., or, if distillation has not been carried to such a high temperature, it may be the same portion plus a gas oil portion that has not been distilled off.

lf desired, the residual portions and/ or whole topped crude can be deasphalted by conventional means.

Cycle stocks or cycle oils refer to product fractions from catalytic cracking units and hydrocracking units which boil above the gasoline boiling range, usually between about 400 and about 850 F. A light cycle oil is a cycle oil boiling generally from about 400 to 650 F. A heavy cycle oil is a cycle oil boiling from about 650 to about 850 F. As in the case with the gas oils, it will be recognized that the term cycle oil is not an extremely precise term, and the boiling ranges given are subject to some variations.

PRIOR ART PROBLEMS (l) General.-Heretofore in a refinery the crude oil has been passed to a distillation unit, normally termed a crude column, and the oil has been fractionated into various cuts, including light gasoline, heavy gasoline, light gas oil, heavy gas oil, and a residual portion boiling so high as to resist vaporization in the crude column. The residual portion has been withdrawn as bottoms and occasionally has been subjected to further processing, for example distillation under vacuum, thermal cracking, steam stripping or coking, to recover additional gas oil distillate fractions.

Heretofore in refinery operations conducted for the purpose of maximizing the conversion of gasoline, 'both the light and the heavy gas oils, but not the aforesaid residual portion, have been used as a feed to a conventional catalytic cracking unit (cat. cracker), although, in some cases, a thermal cracker has been used either alone or in conjunction with the catalytic cracking unit.

In these conventional refinery operations, the gas oil fractions recovered from the crude oil are converted in a the catalytic cracking unit in varying measure to fractions boiling in the gasoline range. Thus, light gas oils can be converted to gasoline at a yield of approximately 40%, while the yield from heavy gas oils is approximately 60%, these and other percentages given herein being on the basis of volumes of gasoline produced per Volume of feed.

(2) Handling of light cycle oil from cat. cracker.- In addition to gasoline, the effluent from the catalytic cracker contains normally gaseous products, as well as heavier portions boiling above gasoline, including light cycle oil, which is highly refractory to further catalytic cracking. Accordingly, every effort is made to work the light cycle oil into various fuel oil products rather' than to recycle the same to the catalytic cracker. On the other hand, the heavy cycle oil portions of the catalytic cracker efuent are much less refractory and are well adapted to be cracked in the catalytic cracker as they are recycled.

The data given in Table I below show the yield of gasoline and of certain other products which can be obtained in the case of a typical renery operation wherein both light and heavy gas oils recovered from a crude oil are fed to a nid catalytic cracker, and wherein, of the cycle oils recovered from the cracker, only those of heavy character are recycled to the unit. The data presented in this table assume a feed to the cracker of 10,000 b./d. of fresh feed made up of about 5500 b./d. of a light gas oil boiling from about 350 to 650 F. and of about 4500 b./d. of a heavy gas oil boiling from about 650 to ll00 F. A heavy cycle oil portion of the cracker effluent boiling above 650 F. (5000 b./d.) is recycled to the cracking zone, thus giving a total feed to the unit of 15,000 b./d.

Table I B./d. Gasoline l 2 5015 C3-C4 oleiins (available for additional alkylation) 1195 C3 and lighter products (equiv. fuel oil yields) '785 Light catalytic cycle oil y2400 1Includes S91 b./d. alkylate produced with all available iCi and an equivalent amount of Cl olefin. pressurized with available nCi to 9.3 lbs. RVP.

2Leaded octane number (F-l-l-S nil. TEL):100.6 (Research Octane Number) after catalytically reforming the nonalkylate gasoline fractions (boiling from about 250 to 400 F.).

It is to be noted from the data of Table I that the light catalytic cycle oil from the catalytic cracker is not recycled. It is found that, when said oil is recycled under conditions wherein approximately 55% thereof is converted to lighter products, the total feed to the catalytic cracker (holding fresh feed constant at 10,000 b./d.) rises from a level of 15,000 b./d. to one of 19,800 b./d. The product yield structure then becomes as follows:

Table Il 13./ d. Gasoline 1 2 5960 C3-C4 olefins (available for additional alkylation) 1380 C3 and lighter products (equiv. liquid yields) 996 Light catalytic cycle oil i100 1Includes 753 b./d. alkylate produced with all .available iCi and an equivalent aioyiint of Ci olefins. Pressurized with available nC to 9.3 lbs.

2Leaded ctnne number (F-'l-l-S ml. TEL):100 .G (Rcsearch Octane Number) after catalytically reforming the nonalkylate gasoline fractions (boiling from about 250 to 400 un),

Comparing the data given in Table I with that of Table Il, it will be seen that the recycling of a substantial portion of the light catalytic cycle oil, though leading to a modest increase in total gasoline production, is economically disadvantageous since it necessitates greatly increasing the capacity of the cracking unit if the fresh feed rate is to be maintained constant.

Inasmuch as the greater portion of the light catalytic cycle oil has its origin in the light gas oil fed to the cracker, it would be desirable if a method were available for effective, eflicient conversion of said gas oil to gasoline without passing the same through a conventional cracking unit. It would also be desirable if a method were available whereby such light catalytic cycle oil, as is produced during cracking, could be converted to gasoline fractions in a similarly eiiicient manner, again without passing said stocks to the catalytic cracking unit.

(3) Handling of residual portion from crude column- As stated above under (1) General, heretofore the residual portion from the crude column has been withdrawn from the system or, occasionally, has been subjected to such further processing steps as vacuum distillation, thermal cracking, steam stripping or coking.

The residual portion has a high asphaltene content, and usually a high metals content. Accordingly, heretofore this portion could not be sent to a conventional catalytic cracker, which is adversely affected both by asphaltenes, which cause an inordinate amount of coking of the Vcracking catalyst, and by metals, which foul the catalyst'.

ln the prior art processing alternatives to which the residual portion has been subjected, that portion has been converted in a large measure to coke or low value black fuel. When it has been sent to a coker, the products are unstable and are poorer feed stocks for subsequent catalytic cracking due to the nonselective nature of lthe coking process, even though the carbon-hydrogen ratio of the hydrocarbons has been improved by the formation of coke. From the standpoint of straight run residual stocks, coking may be competitive with solvent deasphating in the preparation of feed stocks for subsequent catalytic cracking if there is a market for the coke and complete elimination of black fuels and asphalt is desired. Likewise, coking may be competitive with solvent deasphalting since coker gas oils are generally of lower metal and asphaltene content than deasphalted oil and may be included in catalytic cracker feed even though they are inferior 'to virgin stocks. Residual processing schemes, such as vacuum distillation, visbreaking and recycle thermal processing, produce a large percentage of black fuel.

In view of the foregoing, and of major importance to the producer of gasoline and middle distillates, it would be desirable if a practical process were available whereby the residual portion from the crude distillation column could be much more completely utilized than heretofore in the production of gasoline and middle distillates.

(4) Maximizing yield of high fue] value liquid pod ucts per barrel of crude.--Heretofore, partially because of the handling of the aforesaid residual portion, even' very advanced renery process combinations have pro duced considerable quantities of coke and/ or low value liquid products, resulting in a somewhat less than satisfactory yield per barrel of crude of high fuel value liquid products. It would be desirable if a process were provided by means of which this yield could be significantly increased.

(5) Production of heavy naplltlia and light naplztha.- As is well know to those familiar with conventional catalytic cracking and hydroeracking, the major products of such cracking operations are dry gas, butanes, Cgi light naphtha, heavy naphtha and cycle stock (boiling at temperatures higher than about 390 E). In each case, the light naphtha fraction has a relatively high octane number. On the other hand, the heavy naphtha fraction, "particularly that which is obtained by cracking in the presence of hydrogen, has an octane number that is generally several numbers lower. Accordingly,- in order to produce a finished gasoline having a relatively high octane number, it has been the practice tov blend a heavy iiaphtha fraction with a light naphtha fraction, and with butanes in amounts limited by the.

maximum permissibleyaporpressure.' There is, however, a steadily increasing demand for higher octane gasolines (about 95 and higher). Those skilled in the art will readily appreciate that such octane requirements cannot be met by the aforementioned conventional blending operations. Accordingly, the relatively low octane heavy naphtha fraction has been subjected to reforming operations. As the increasing demand for higher octane gasolines must be satisfied by reforming, instead of by blending, there is a greater demand for heavy naphtha fractions that can be reformed and a correspondingly lesser demand for light naphtha fractions that can be used for blending purposes. The light naphtha fraction usually is not subjected to a reforming operation because it produces excessive amounts of dry gas, coke, etc. The yield of gasoline obtained by reforming light naph'tha therefore is prohibitively small. Accordingly, it is very desirable that there be provided a process that will produce greater amounts of heavy naphtha and lesser amounts of light naphtha.

(6) Production of gasoline product with satisfactory road octane number.-Heretofore, conventional ca'talytic cracking processes have been used to convert a portion of the crude feed to heavy gasoline, which then has been reformed to increase its octane number. However, the resulting reformate has not had a satisfactory road octane number without additional blending with an alkylate produced from isobutane and olelins, and sufficient quantities of alkylate have not been obtainable in the same process to meet the blending requirements of the catalytic reformate. By definition, road octane is A(F-l)-l-B(F2)}-C, Where F-l and F-2 are the conventional octane ratings, and Where A, B and C are constants, and where A is about the same order of magnitude as B. Also by definition, sensitivity is (F-l)-(F2). Accordingly, a high sensitivity adversely affects road octane number, and conversely, a low sensitivity is helpful to road octane number.

Generally, a catalytically cracked heavy gasoline has a sensitivity of around 1l to 13, and generally the reformate resulting from the reforming of this heavy gasoline has a sensitivity of around to 11.5. Accordingly, to obtain a satisfactory road octane number, the reformate must be blended with alkylate, which has a sensitivity of around 1. Where sufficient quantities of alkylate have not been available from the process, the deficiency has had to be made up with alkylate from another source.

In view of the foregoing, it would be desirable if an integrated process were provided which would produce sufhcient quantities of alkylate and of products having a similarly low sensitivity, to satisfy the blending requirements of the catalytic reformate produced in the process.

(7) Flexibility in ratio of light gas oil product to gasoline product-Prior art processes generally are capable of producing fair yields of light gas oil from the crude barrel, or fair yields of gasoline. However, there has been a marked rigidity in the capabilities of prior art processes to accomplish production of these two products in a widely Varying ratio in response to demand changes caused by seasonal fluctuations and other factors. It would be desirable if a process Were available to accomplish such production.

OBI ECTS The objects of the present invention include solution to the foregoing problems; accordingly, the objects include the following:

(1) It is a general object of this invention to provide an integrated refinery process wherein, in addition to the gasoline fractions, all of the gas oil fractions distilled yfrom a crude hydrocarbon oil source and the remaining residual portion can be eficiently converted to gasoline fractions of high octane rating.

(2) Another object of the present invention is to provide a process of this character which integrates the action of a conventional crackingl unit, operating efciently with feed stocks, including heavy gas oils, which are readily yconverted therein, with that of a hydrocracking unit which receives, and is capable of readily converting to gasoline, the light gas oil fractions from the crude source and light cycle `stocks from the crack- Iing unit.

(3) Another object of the invention is to integrate into a process of this character a particular means for handling said residual portion so that it may be substantially completely utilized in the integrated process for ultimate conversion to gasoline, instead of all or a su` stantial portion thereof being directed to less valuable uses as has been necessary heretofore.

(4) Another object of the present invention is to provide a process of the foregoing character with which may be `maximized lthe total liquid yield per barrel of crude of high fuel value liquid products.

(5) Another object of the present invention is to provide a process of the foregoing character with which large amounts of heavy naphtha suitable lfor catalytic -reforming may be produced and in which at least a portion of the desired heavy naphtha fraction is produced from a residual portion of a crude petroleum feed stock.

(6) Ano-ther object of the present invention is to provide a process `capa-ble of producing a catalytic reformate yand also sufficient amounts of alkylate and other products having -loW sensitivity to satisfy the 'blending requirements of the catalytic reformate.

(7) Still another object of the present invention is to lprovide a process with which the ratio of light gas oil product to lgasoline product may be Varied over a wide range.

STATEMENT OF INVENTION in accordance with a specific embodiment of the present invention, there is provided a process for converting substantially all of a crude hydrocarbon feed to fuel values in an exceptionally high liquid yield, which cornprises separating said crude feed into `fractions including light gasoline, heavy gasoline, gas oil and residuum fractions, `converting at least a portion of said gas oil to gasoline in a hydrocracking zone, converting at least la portion of said gas oil to gasoline in a catalytic cracking Zone, and converting at least a portion of said residuum into gasoline `in a catalytic cracking zone operating `at low conversion.

`In a more specific embodiment of the present invention, there is provided a process for converting substantially all of a crude hydrocarbon feed to fuel values in an exceptionally high liquid yield, which comprises separating said crude feed into Ifractions including light gasoline, heavy gasoline, light gas oil, heavy gas oil and residuum fractions, recovering said light gasoline as a net product, converting said heavy gasoline in a catalytic reformer to gasoline of higher octane number, converting a substantial portion of said light gas oil in a hydrocracking zone to gasoline, converting a substantial portion of said heavy gas Oil in a catalytic `cracking zone to gasoline and cycle oil, and converting a substantial portion of said residuum, in a catalytic cracking zone operating at low conversion, to gasoline and cycle oil.

In a still more specific embodiment of the present invention, there is provided a process for converting substantially all of a crude hydrocarbon feed 4to fuel values in an exceptionally high liquid yield, in an integrated `system of refinery units including a hydrocracking zone, a conventional catalytic cracking zone, a catalytic reforming zone, and a second catalytic cracking zone oper- [ating -a-t low conversion, which comprises separating said crude feed into fractions including light gasoline, heavy gasoline, light gas oil, heavy ygas oil and residuum fractions, recovering said light gasoline as a net product, passing at least a portion of said heavy gasoline to said reforming zone, passing at least a portion of said light Z gas oil to said hydrocracking zone, passing at least a portion of said heavy gas oil to said conventional catalytic cracking zone, passing at least a portion of said residuurn to said second catalytic cracking zone, recover- Ying a light gasoline `from said second catalytic cracking zone, passing a heavy gasoline from said second catalytic cracking zone to said reforming zone, passing a light cycle oil from said second `catalytic cracking zone to said hydrocracking Zone, passing a heavy cycle oil from said second catalytic cracking zone to said conventional catalytic cracking zone, passing a light cycle o-il from said conventional catalytic cracking zone to said hydrocracking zone, passing a heavy gasoline from said conventional catalytic cracking zone to said reforming zone, passing a heavy gasoline `from said hydrocracking zone to said reforming Zone, recovering a light gasoline from said hydrocracking zone, and recovering a high octane gasoline refoirnate from said reforming zone.

DRAWING The present linvention will best be understood, and further objects and advantages thereof will be apparent, from the 'following detailed description, when read in connection with the accompanying drawing, which is a simplified ow diagram illustrating a group of refinery units and interconnecting flow paths suitable for carrying out the process of the invention.

PROCESS UNITS AND OPERATING CONDITIONS, GENERAL The process of the present invention may include, in various combinations, distillation zones, a conventional catalytic cracking zone, a second catalytic cracking zone operating at low conversion, a hydrocracking zone, a catalytic reforming zone, an alkylation zone, and a deasphalting zone.

Suitable catalysts and operating conditions for these various zones are described immediately below.

DISTILLATION ZONES The necessary distillation zones and operating conditions thereof for the process of the `present invention are conventional, from the crude column to the miscellaneous distillation zones operating in conjunction with the various conversion units.

CATALYTIC CRACKING ZONES The catalytic cracking zones in the process of the present invention are conventional catalytic cracking zones operating with a conventional catalytic cracking catalyst, for example silica-alumina or silica-magnesia, at conventional catalytic cracking conditions, for example, a temperature ot' about from 850 to 1000" F. for one Zone and 950 to l000 F. for the other zone, ask discussed below, and a pressure of .about from 10 to 30 p.s.i.g. As will be discussed hereinafter, one of the catalytic cracking zones operates at a low conversion, for example less than 35% per-pass conversion, and the other catalytic cracking zone operates at conventional conversion levels, for example S to 85% per-pass conversion. The low conversion catalytic cracking zone is operated at a temperature above 950 F.

Desirably, the low conversion catalytic cnacking zone is a transfer line catalytic cracker. The catalytic cracking zone operating lat conventional conversion levels may be either of the moving bed or liuid type.

HYDROCRACKING ZONE While the invention can be practiced with utility in connection with feeds to the hydrocracking zone containing relatively large quantities of nitrogen, the operation becomes much more economical with stocks containing less than 200 ppm., preferably below 100 ppm., and, much more preferably, below ppm., of nitrogen. A reduction in feed nitrogen level permits the hydrocracking reaction to be conducted at lower temperatures than with feeds containing relatively large amounts of nitrogen compounds. Therefore, in the case of hydrocracking zone feed stocks which are not inherently low in nitrogen, acceptable levels can be reached by pretreating the feed to the catalytic cracking unit by a nitrogen cornpound extraction process, or "by contacting either the catalytic cracking unit feed or, preferably, the particular feed to the hydrocracking process, with hydrogen in the presence of a suitable catalyst at elevated temperatures and pressures to remove nitrogen compounds therefrom. A particularly effective catalyst for removing nitrogen by hydrogenation is one wherein a coprecipitated molybdena-alumina material (eg, one prepared in accordance with the disclosure of U.S. Patent 2,432,286 to Claussen et al. or U.S. Patent 2,697,006 to Sieg) is combined with cobalt oxide, the final catalyst having a metals content equivalent to about 2% cobalt and 7% molybdenum. Representative processing conditions for removing nitrogen with this catalyst are an LHSV of 1 to 3, 700 to 800 F., 200 to 2500 and 1000 to 15,000 s.c.f. of hydrogen per barrel of feed stock.

When nitrogen removal is effected by hydroning, the resulting effluent is treated, in accordance with methods presently known in the art, so as to remove ammonia and some hydrogen sulfide which may be present. A preferred removal method involves injecting Water into the total efliuent from the hydroining unit and then passing the resulting mixture into a high pressure separator operating under such conditions of temperature and pressure (for example F. and 950 psig.) that a gaseous overhead is removed that is predominantly hydrogen, but which normally contains some hydrogen sulfide and light hydrocarbons. This overhead (following a clean-up treatment to remove any nitrogen and sulfur-containing' compounds, if desired) can be recycled to the hydroning unit along with make-up hydrogen. Two liquid phases are formed in the separator, an upper hydrocarbon phase and a lower aqueous phase which contains essentially .all of the ammonia present and some hydrogen sulde in the form of ammonium sulfide. The aqueous phase is removed from the system and discarded.

The hydrocarbon layer is then preferably passed into a stripper or distillation column from which any remaining hydrogen sulfide, ammonia and Water are removed overhead. The stream may also be freed of any light hydrocarbon fractions (boiling in the gasoline range or below) formed as a result of hydrocracking reactions taking place over' the hydroiining catalyst.

The portion of the denitried feed to be hydro-cracked, along with from about 1500 to 30,000, and preferably from about 3000 to 15,000, standard cubic feet (scf.) of hydrogen per barrel of total reaction feed, is passed into the hydrocracking zone at a liquid hourly space velocity (LHSV) of from about 0.2 to l5, tand preferably from about 0.4 to 3.0, and intimately contacted with the catalyst.

The contacting step is conducted under a pressure of at least 500 p.s.i.g., and preferably from about S00 to 3000 p.s.i.g. The temperature is preferably maintained in the nange of from about 400 to 750 F., because at temperatures above about 750 to 800 F. the amount of gasoline product lost to the less desirable C3 and lighter materials rapidly increases, thus lowen'ng the motor fuel yield. For example, it has been found that the amount of methane produced at 800 F. per unit of converted product is approximately sixteen times as great `as that formed at 700 F., and four times as great as that produced at 750 F. At higher temperatures, the situation becomes much Worse. Accordingly, resort is normally had to temperatures above about 750 F. up to about 850 F. only in the last stages of the catalyst on stream period when it is desired to maintain relatively high activity at the expense of higher light gas losses or, in the case when the relatively high nitrogen-containing feeds are processed. Further, operations at tempenatures above about 750 F. and at low or moderate pressures induce a relatively rapid decrease in the activity of the catalyst as reflected by reduced per-pass conversion levels. Thus, when operating at 375 F. and at a relatively low pressure, for example 1500 psig., on la hydrofined light cycle oil feed, regeneration of cobalt-molybdenum on silica-alumina catalyst is required in most instances after on-stream periods of one day or less, which contrasts with ori-stream periods of 100 to 300 or more hours at good activity as temperatures are maintained below about 825 F. With operation at 800 F. and higher with the `same and similar feeds, but with nickel sulfide on silicaalumina, regeneration is required after on-stream periods of a few hundred hours or less, compared with Ioperation below 700 F., with which can be obtained on-stream periods of several thousand hours without regeneration. In the present process, it is recommended that the reaction be conducted at an anitial on-stream temperature fro-m about 550 to 650 F., with a progressive increase to about 750 to 850 F. so as to maintain catalyst activity at a controlled level. The initial and terminal temperatures will vary, with character of feed and catalyst, within the 4overall range specified above.

The catalyst employed in the hydrocracking zone is one wherein a material having hydrogenating-dehydrogenating activity is deposited or otherwise disposed on an active cracking catalyst support. rThe cracking component may comprise any one or more `of such acidic materials as silica-alumina, silica-magnesia, silica-aluminazirconia composites, alumina-boria, fluorided composities, and the like, as well yas various acid-treated clays and similar materials. The hydrogenating-dehydrogenating components of the catalyst can be selected from any one or more of the various groups VI, VII and VIII metals, as well as the oxides and sullides thereof, alone or together with promoters and stabilizers that may have by themselves small catalytic effect, representative materials being the oxides and suliides of molybdenum, tungsten, vanadium, chromium and the like, as well as of metals such as iron, nickel, cobalt and platinum. If desired, more than one hydrogenating-dehydrogenating component can be present, and good results may be obtained with catalysts containing composites of two or more of the oxides of molybdenum, cobalt, chromium, tin and zinc, and with mixtures of said oxides with uorine. The amount of the hydogenating-dehydrogenating component present can be varied within relatively wide limits of from about 0.5 to 30% based on the weight of the entire catalyst.

Exemplary catalysts having satisfactory characteristics as aforesaid include those containing: (a) about 1 to 12% molybdenum oxide, (b) a mixture of from l to 12% molybdenum oxide and from 0.1 to cobalt oxide, (c) mixtures of from about 0.5 to 10% each of cobalt oxide and chromium oxide, (d) 0.1 to 10% nickel, nickel oxide or nickel sulfide, (e) 0.1 to 10% cobalt, cobalt oxide or cobalt sulfide, (f) mixtures of from 0.1 to 10% each of nickel and cobalt, as metal, oxide or sulfide, in each case the said hydrogenating-dehydrogenating component being deposited on an active cracking support comprising silica-alumina beads having a silica content of about 70 to 99%. Thus, the molybdenum oxide catalyst can be prepared readily by soaking the beads in a solution of ammonium molybdate, drying the catalyst for 24 hours at 220 F., and then calcining the dried material for 10 hours at 1000 F. If cobalt oxide is also to be present, the calcined beads can then be similarly treated with a solution of a cobalt compound, whereupon the catalyst is again dried and calcined. Nickel sulfide and/ or cobalt sulfide are especially suitable. The entire preferred catalyst composite and conditions of Scott U.S. Patent 2,944,006 will be especially suitable in the process of the present invention. Under favorable operating conditions, the hydro-cracking catalyst will maintain high l0 activity over periods of 50 to 300 or more hours. The activity of the used catalyst can then be increased, if desired, by a conventional regeneration treatment involving burning olf catalyst contaminants with an oxigen-containing gas.

The light gasoline produced in the hydrocracking zone has a very low sensitivity, and is well suited for blending with catalytic reformate having an unsatisfactory road octane number, to produce a blend of higher road octane number than the reformate.

REFORMING ZONE The reforming zone is operated at conventional reforming conditions including temperatures of about 700 to about 1000" F., preferably between about 725 land about 950 F. The LHSV will vary between about 0.1 and about 10, preferably between about 0.5 and 4. The hydrogen pressure will vary between about 100 p.s.i.g. and about 1000 psig., preferably between about 350 and 750 p.s.i.g. The molar ratio of hydrogen to hydrocarbon charge will vary between about 1 and 20, preferably between about 4 and 12.

The reforming operation is carried out in the presence of hydrogen and a suitable reforming catalyst. Such catalysts include the metals and compounds such as the oxides and/or sulfides of the metals `of the left-hand column of Groups VI and VH1 of the Periodic Table of the Elements. The catalyst compounds can be used alone or on a suitable support. Catalysts that comprise platinum or palladium metal deposited on supports such as silica-alumina, silica-Zirconia, alumina-boria, aluminahalogen-activated alumina and the like are particularly suitable.

In the reforming process, there is net production of hydrogen. This net production of hydrogen may be used all or in part to supply the hydrogen requirements for the hydrocracking zone.

ALKYLATION ZON E The alkylation zone is operated with a conventional alkylation catalyst, for example H2804 or HF, at conventional alkylation conditions.

With an H2504 catalyst, the operating temperature may be from about to 55 F. using make-up acid of approximately 99% concentration. With propylene present, the temperature should be kept at least as high as F.

With HF, the temperature can be from about 40 to about 100 F.; with this catalyst, temperature effects are not so marked `as with H2804, and because of internal catalyst regeneration acid consumption is not such a problem as it is with H2804.

With either catalyst, the reaction rates will be lower with lower operating temperatures, and with either catalyst a pressure will be used that is suliicient to maintain a liquid phase at the operating temperature. A higher pressure will be required with HF.

Generally, it is desirable to have an isobutane concentration of about from to 75% in the reaction mix; with a higher isobutane concentration, the product will have a higher octane number.

DEASPHALTIZING ZONE Where a deasphaltizing zone is used, i-t may be any conventional deasphaltizing zone, for example a propane deasphalting zone, operating under conventional deasphalting conditions.

DETAILED DESCRIPTION A best mode for carrying out the process of the present invention may be determined by reference to the appended drawing which is a diagrammatic illustration of a group of interrelated refinery units and flow paths suitable for use in practicing the process. For purposes of clarity and because their location and use will be readily apparent to those skilled in the art, various pieces of conventional equipment, such as heaters and pumps, have been omitted from the drawing. The following detailed description will indicate how the process of the present invention may be operated to maximize gasoline production. With this description as a guide, those skilled in the art will understand that the operation easily may be varied to maximize middle distillation production, mainly by simply taking as products the middle distillate range materials, such as light cycle oils, from the various process units, rather than further processing them to produce gasoline, and by adjusting the recycle cut points to the various process units as desired.

A crude oil feed is supplied through line 1 to crude column 2. Desirably, the crude oil feed is a parafiinic crude, for example a Mid-Continent, Middle East, Canadian, West Texas or East Texas crude. In crude column 2, the crude is stripped of a light gasoline fraction which is removed as a product through line 3, a heavy gasoline fraction which is passed through line 4 to catalytic reformer 5, a light gas oil fraction which is passed through line :to hydrocracking zone 11, and a heavy gas oil fraction which may boil from about 550 to 800 F., which may be passed through line 12 to conventional catalytic cracking zone 13, as shown, or to hydrocracking zone 11. Metal contaminants in the crude oil feed tend to concentrate in the higher boiling portions thereof. Accordingly, crude column 2 is operated to limit to about 800 F. the end point of the heavy straight run gas oil in line 12 going to catalytic cracker 13, contamination of the catalyst in catalytic cracker 13 by metals will be less than if the heavy straight run gas` oil had a higher end point, and an adequate supply of heavy straight run gas oil for catalytic cracker 13 still will be possible. Catalytic cracking zone 13 may be any conven-tional catalytic cracking unit, such as a fixed-bed catalytic cracko er of the Houdry commercial type, a Thermofor catalytic cracker, or a fluid-type moving-bed catalytic cracker. If desired, a portion of the light gas oil from line 10 may be withdrawn through line 10A, for example for use as a No. 2 oil. With the crude thus stripped to about 800 F., a residuum comprising about of the original crude feed remains, and is passed through line 14 to solvent deasphalting zone 15, which may be a conventional propane deasphalting zone, and thence through line 16 to transfer line catalyticc racker 17. It is desirable to include in the residual stock in line 16 to transfer line catalytic cracker 17 all portions of the crude feed boiling as low as 800 F.; if the residual stock includes only higher boiling material, for example only materials boiling above 900 F., there will be a great er percentage of carbon residue in the stock. This residue will appear as increased production of coke on the catalyst of zone 13, and that unit accordingly will convert a decreased percentage of i-ts feed to liquid products.

Transfer line catalytic cracker 17 is operated at a low conversion, for example yto 35% per pass, with a low activity silica-alumina catalyst. A portion of the catalyst in zone 17 may be periodically or continuously discarded and replaced with regenerated catalyst withdrawn from the conventional catalytic cracker 13; the catalyst Withdrawn from catalytic cracker 13 may be replaced with fresh cracking catalyst to maintain a high cracking activity in catalytic cracker 13. Alternatively, the make-up of low activity catalyst required in zone 17 may be supplied from external sources. In catalytic cracker 17 residuum is cracked or decarbonized in concurrent flow upwardly with catalyst circulating through line 20, catalyst and oil separator 21, line 22, conventional catalyst regeneration zone 23 where coke is burned from the catalyst, and line 24. Periodically, a portion of the regenerated catalyst in transfer line catalytic cracker 17 may be withdrawn and treated in a conventional demetalation zone, not shown, .to remove from the catalyst the accumulation of metals resulting from processing of the feed stock in zone 17 lthe thus treated catalyst then may be returned to catalytic cracker 17. Even the heavy cycle oil, the heaviest product from catalytic cracker 17, does not contain any significant amount of asphaltenes or metals, both of which are removed from the feed and deposited on the catalyst of this unit. Catalytic cracker 17 does not produce low value residual black fuel from the residual feed, but produces only distillate products. The reactor temperature and pressure may be conventional, for example 1000 F. and 15 psig., respectively, but must be above 950 F. The reactor may be operated with a catalyst to oil ratio of about 8.0. Under these conditions the low conversion may be obtained by operating the reactor with a space velocity above about 25.

In catalyst and oil separator 21, the catalyst is separated from the oil product and passed through line 22 to regenerator 23, and the oil product is passed through line 25 to distillation column 26, where it is separated into various fractions. A C3: and C4: fraction is passed through line 27 to alkylation zone 2S. A light gasoline fraction is withdrawn as a product through lines 29 and 30. A heavy gasoline fraction may be passed through lines and 4 to catalytic reformer 5. However, because of the high octane number of this fraction, it may be desired to withdraw all or a portion of it as a product, instead of reforming it, particularly where it is not necessary to hydrofine it to remove such contaminants as sulfur, nitrogen, diolens and other gum-forming precursors. Hydrotining will tend to reduce the octane number and necessitate reforming. A light cycle oil fraction, which would be refractory to catalytic cracking, is passed through line to hydrocracker 11. A heavy cycle oil may be passed through lines 41 and 41A, and then may either be passed through line 42 to catalytic cracker 13, or may be passed through line 43 to conventional extraction zone 44. In the latter case, extraction zone 44, which may be operated with a conventional solvent such as SO2, phenol or furfural, operates to separate the heavy cycle oil entering through line 43 into `an aromatic extract fraction and a raffinate fraction. The extract fraction is passed through line 45 to hydrocracker 11, and the raftinate fraction is passed through line 46 to catalytic cracker 13. Alternatively, instead of passing the heavy cycle oil from distillation column 26 through lines 41 and 41A, it may be passed through line 41B to hydrocracker 1l.

Catalytic cracker t3 accordingly may be supplied with heavy straight run gas oil through line 12 and possibly with either heavy cycle oil through line 42 or a paraliinic raffinate through line 46. Catalytic cracker 13 also may be supplied with a heavy cracked cycle oil through lines and 4Z, and may be supplied with a 400 F.-| hydrocracked cycle oil through line 56. i

Catalytic cracker 13 is operated under conventional catalytic cracking conditions -at a higher conversion, for example 75 to 80% per pass. The reaction products from catalytic cracker 13 are passed through line 57 to distillation column S8, where they are separated into various fractions. A C3: and C4: fraction is passed through line 59 to alkylation zone 28. A light cycle oil fraction is passed through line 60 to hydrocracking zone 11. A heavy gasoline fraction is passed through lines and 4 to catalytic reformer S. However, because of the high octane number of this fraction, it may be desired -to withdraw all or a portion of it as a product instead of reforming it, particularly Where it is not necessary to hydroiine it to remove such contaminants as sulfur, nitrogen, dioleins and other gum-forming precursors. Hydroning will tend to reduce .the octane number and necessitate reforming. A heavy cycle oil is recycled through lines 55 ,and 42 to catalytic cracker 13.

Catalytic cracking catalyst is cycled through line 70, catalyst regeneration zone '71, where the catalyst is regenerated in conventional manner, and line 72.

Hydrocracker 11 is supplied with a light gas oil -through line 10, a light cycle oil through line 40, an aromatic extract through line 45 if desired, a 400 F.`{ cycle oil through line 73, if desired, a heavy cycle oil through line 41B if desired; all or a portion of the heavy straight run gas oil in line 12 may be sent to hydrocracker 11 instead of catalytic cracker 13 if desired. The reaction products from hydrocracker 11 are passed through line 74 to distillation column 75 Where they are separated into various fractions.

From distillation column 7S, an isohutane frac-tion is passed through line 80 to alkylation zone 28. A light gasoline fraction is recovered through lines 81 and 30 as a product. A heavy gasoline fraction is passed through lines 82, 65 and 4 to catalytic reformer 5. A 400 F.'+ cycle oil is passed through line 83, and thence through either line 73 to hydrocracker 11 or line 56 to catalytic cracker 13, as desired.

C3: and C4: fractions and isobutane from the two catalytic cvracking zones enter alkylation Zone 28 through lines 27 and 59, respectively, and these C3: and C4: fractions are alkylated there with said isohutane, with isohutane from hydrocracking zone 11 which enters alkylation zone 28 through line 80, and, if desired, with isobutane from reforming zone 5. A high octane product alkylate is withdrawn from alkylation zone 2S through line S4.

In catalytic reformer 5, a heavy straight run gasoline fraction from crude column 2, a heavy gasoline fraction from catalytic cracker 17, the heavy gasoline fraction from hydrocracker 11, and the heavy gasoline fraction from catalytic cracker 13 are reformed under conventional reforming conditions to produce a high octane reformate which is withdrawn from catalytic reformer through line 85.

EXAMPLES The following'example illustrates the processing of 50,000 barrels per day of a Middle East crude oil to produce a maximum yield of gasoline, using a conventional refinery arrangement including delayed coking, catalytic cracking, C4 alkylation and catalytic reforming, compared With the processing of the same daily quanti-ty of the same feed in a refinery arrangement in accordance with the present invention, also to produce a maximum yield of gasoline.

EXAMPLE 1 The following example illustrates the processing of the some daily quantity of the same feed `as in Example 1, to produce a maximum yield of middle distillates, using the same conventional refinery arrangement as in Example 1, compared with the processing of the same daily quantity of the same feed in a refinery arrangement in accordance with the present invention, also to produce a maximum yield of middle distillates.

EXAMPLE 2 [Maximizing Middle Distillates] Conventional Present Invention Processing Arrangement (Same as v (Same as Example 1) Example l) Refinery Feed, Refinery Input, b./d.:

34.5 API Middle East Crude 50, 000 50, 000 i-Butane 1, 000

Total 51, 000 50, 000

Products:

Motor Gasoline, 10 lb. RVP 1 21, 800 21,900 F-l-i-B 95.2 95. 2 F-2-l-3- 89. 0 90. 0 Road A 95.0 96. 0

Middle Distillate 20, 250 31, 500

Middle Distillate/ Gasoline 0.9 1. 45

l Reid Vapor Pressure.

From the foregoing examples, it will be seen that compared with the conventional refinery arrangement, the process of the present invention results in: (a) higher yield of more valuable liquid products, i.e., 100 volume percent of refinery input comprising isobutane and motor gasoline, compared to volume percent in the conventional refinery; (b) complete elimination of refractory,

cracked gas oils of low quality; (c) complete elimination of residual fuel oil, including low quality coke; (d) maximizing production of heavy naphtha for upgrading into high octane aromatic reformates; (e) improvement in gasoline poolvroad octane 'by substituting isoparafiinic gasoline for olefinic compounds; and (f) ability to vary middle distillate/ gasoline ratio between 0 and 1.45 cornpared to 0.23 and 0.9 for conventional refinery processing.

[Maximizing gasoline] Conventional Present Invention Case A Case B Processing Arrangement Delayed Coking Catalytic Cracking C4 Alkylation Catalytic Reforming Transfer Line Catalytic Cracking Hydrocracking C4 Allrylation Catalytic Cracking Catalytic Reforming 1 EFO (Equivalent Fuel Oil) is amount of 10 API bunker fuel that would have equivalent heating value in B.t.u., assuming that 1 barrel of said fuel oil has a heating value of 6.3M Btu.

2 Reid Vapor Pressure. 3 Tons per day.

l SUMMARY From the foregoing detailed description, it may be seen that the process of the present invention operates in an extremely eicient manner to utilize the entire spectrum of the crude oil feed in making only liquid products of high fuel values in high liquid yields, with -a minimum installation of expensive renery equipment. Such an integrated refinery, capable of converting all of the crude feed to liquid productsl of high value with minimum waste and losses, long has been a goal desired by refiners.

It will be apparent that the net liquid product streams shown easily may be utilized by the refinery as fuels, in blending operations, or subjected to further processing, as desired. Any conventional blending operations may be used in connection with the process of the invention. For example, in a conventional gasoline blending zone, not shown in the drawing, the high octane reformate from catalytic reformer 5 may be blended with butanes andrwith C5+ light naphtha, in suitabie proportions to produce a finished gasoline of the desired Reid vapor pressure and octane number.

Although only specific modes of operation of the process of the present invention have been described, numerons variations could be made in those modes without departing from the spirit of the invention, and all such variations that fall 'within the scope of the appended claims are intended to be embraced thereby.

Iclaim:

1. In a process which comprises separating a crude hydrocarbon feed into fractions including gas oil and residual fractions, converting at least a portion of said gas oil to gasoline in a hydrocracking zone, converting at least a portion of said gas oil to gasoline in a catalytic cracking zone, and converting at least a portion of said residual fraction to gasoline in a catalytic cracking zone,

the improvement which comprises converting substantially all of said hydrocarbon feed to fuel values in an exceptionally high liquid yield by converting at least a portion of said gas' oil to gasoline in a iirst catalytic cracking zone operated at conventional catalytic cracking zone conversion levels, converting said residual fraction -to gasoline in a second catalytic cracking zone operated at a low per-pass conversion not exceeding about a temperature above 950 F., and a sufficiently high space velocity in conjunction with said temperature to produce said low per-pass conversion, and recovering from said hydrocracking zone, said irst catalytic crackl5 ing zone, and said second catalytic cracking zone gasoline and cycle oil products in high liquid yields.

2. A process as in claim 1 wherein said second catalytic cracking zone is operated at a space velocity above about 20 v./v./hr.

3. In a process which comprises separating a crude hydrocarbon feed into fractions including gas oil and residual fractions, converting at least a portion of said gas oil to gasoline in a hydrocracking zone, converting at least a portion of said gas oil to gasoline in a catalytic cracking zone, and converting at least a portion of said residual fraction to gasoline in a catalytic cracking zone, the improvement which comprises separating said crude feed into fractions including light gasoline, heavy gasoline, light gas oil, heavy gas oil and residual fractions, recovering said light gasoline as a net product, converting said heavy gasoline in a catalytic reformer to gasoline of higher octane number, converting a .substantial portion of said light gas oil in a hydrocracking zone to gasoline, converting a substantial portion of said heavy gas oil in a cracking zone operated at conventional conversion conditions to gasoline and cycie oil, converting a substantial portion of said residual fraction to gasoline and cycle oil in a second catalytic cracking zone operated at a temperature above 950o F. and at a low per-pass conversion, recovering a high octane gasoline product from said catalytic reformer, recovering a gasoline product from said hydrocracking zone, and recovering gasoline products from said catalytic cracking zones.

4. A process as in claim 3 wherein said second catalytic cracking zone is operated at a space velocity above about 20 v./v./hr.

References Cited by the Examiner UNITED STATES PATENTS 2,360,622 10/44 Roetheli 208-80 2,528,586 11/50 Ford 208-80 2,998,380 8/61 McHenry et al. 208-80 3,008,895 ll/6l Hansford et al 208-112 3,072,560 l/ 63 Paterson et al 208-55 OTHER REFERENCES The Effects of Variables in Catalytic Cracking, Oblad et al., pages to 188, vol. 2, Chemistry of Petroleum Hydrocarbons, Reinhold Pub. Corp., New York, 1955.

ALPHONSO D. SULLIVAN, Primary Examiner.

JOSEPH R. LIBERMAN, Examiner. 

1. IN A PROCESS WHICH COMPRISES SEPARATING A CRUDE HYDROCARBON FEED INTO FRACTION INCLUDING GAS OIL AND RESIDUAL FRACTIONS, CONVERTING AT LEAST A PORTION OF SAID GAS OIL TO GASOLINE IN A HYDROCRACKING ZONE, CONVERTING AT LEAST A PORTION OF SAID GAS OIL TO GASOLINE IN A CATALYTIC CRACKING ZONE, AND CONVERTING AT LEAST A PORTION OF SAID RESIDUAL FRACTION TO GASOLINE IN A CATALYTIC CRACKING ZONE, THE IMPROVEMENT WHICH COMPRISES CONVERTING SUBSTANTIALLY ALL OF SAID HYDROCARBON FEED TO FUEL VALUES IN AN EXCEPTIONALLY HIGH LIQUID YIELD CONVERTING AT LEAST A PORTION OF SAID GAS OIL TO GASOLINE IN A FRIST CATALYTIC CRACKING ZONE OPERATED AT CONVENTIONAL CATALLYTIC CRACKING ZONE CONVERSION LEVELS, CONVERTING SAID RESIDUAL FRACTION TO GASOLINE IN A SECOND CATALYTIC CRACKING ZONE OPERATED AT A LOW PER-PASS CONVERSION NOT EXCEEDING ABOUT 35%, A TEMPERATURE ABOVE 950*F., AND A SUFFICIENTLY HIGH SPACE VELOCITY IN CONJUCTION WITH SAID TEMPERATURE TO PRODUCE SAID LOW PER-PASS CONVERSION, AND RECOVERING FROM SAID HYDROCRACKING ZONE, SAID FIRST CATALYTIC CRACKING ZONE, AND SAID SECOND CATALYTIC CRACKING ZONE GASOLINE AND CYCLE OIL PRODUCTS IN HIGH LIQUID YIELDS. 